Process for the dissociation of MTBE

ABSTRACT

Isobutene is prepared by a process in which a) an MTBE-containing stream I is separated by distillation into an MTBE-containing overhead stream II and a bottom stream III which comprises compounds having boiling points higher than that of MTBE; and b) the MTBE present in the overhead stream II is dissociated over a catalyst to give a dissociation product IV; wherein the stream I has a proportion of 2-methoxybutane (MSBE) of greater than 1000 ppm by mass, based on MTBE, and wherein the separation by distillation in step a) and/or the dissociation in step b) is carried out so that the dissociation product IV has a concentration of less than 1000 ppm by mass of linear butenes, based on a C 4 -olefin fraction.

BACKGROUND OF THE INVENTION

1. Field of the Invention

The present invention relates to a process for preparing isobutene bydissociation of methyl tert-butyl ether (MTBE).

2. Description of the Related Art

Isobutene is a starting material for the production of many products,e.g. for the production of butyl rubber, polyisobutylene, isobuteneoligomers, branched C₅-aldehydes, C₅-carboxylic acids, C₅-alcohols andC₅-olefins. It is also used as an alkylating agent, in particular forthe synthesis of tert-butylaromatics, and as intermediate for theproduction of peroxides. In addition, isobutene can be used as precursorfor the preparation of methacrylic acid and its esters.

In industrial streams, isobutene frequently occurs together withsaturated and unsaturated C₄-hydrocarbons. Owing to the small boilingpoint difference or the very low separation factor between isobutene and1-butene in a distillation, isobutene cannot be separated economicallyfrom these mixtures. Isobutene is therefore usually isolated fromindustrial hydrocarbon mixtures by converting isobutene into aderivative which can easily be separated off from the remaininghydrocarbon mixture and redissociating the isolated derivative to formisobutene and the derivatizing agent.

Isobutene is usually separated off from C₄ fractions, for example the C₄fraction from a steamcracker, as follows. After removal of the majorpart of the multiply unsaturated hydrocarbons, mainly butadiene, byextraction/extractive distillation or selective hydrogenation to linearbutenes, the remaining mixture (raffinate I or hydrogenated cracking C₄)is reacted with alcohol or water. The use of methanol results information of methyl tert-butyl ether (MTBE) from isobutene and the useof water results in tert-butanol (TBA). After they have been separatedoff, both products can be dissociated to form isobutene in a reversal oftheir formation.

MTBE is cheaper than TBA because the reaction of isobutene-containinghydrocarbons with methanol is easier than with water and MTBE isproduced in large quantities as a component of 4-stroke fuels. Theproduction of isobutene from MTBE is therefore potentially moreeconomical than from TBA if a similarly good process were to beavailable for the dissociation of MTBE as for the dissociation of TBA.

The dissociation of MTBE can be carried out in the liquid phase, thegas/liquid phase or the gas phase in the presence of acid catalysts.Regardless of the phase in which the dissociation is carried out,by-products are formed, although to differing extents. For example,undesirable C₈ and C₁₂ components can be formed by acid-catalyseddimerization or oligomerization of the isobutene formed in thedissociation. The undesirable C₈ components are mainly2,4,4-trimethyl-1-pentene and 2,4,4-trimethyl-2-pentene. In addition,part of the methanol formed in the dissociation can be converted intodimethyl ether.

For economical reasons (low price; good availability), the preparationof isobutene by dissociation of MTBE is not carried out usinghigh-purity MTBE, for example MTBE/S from Oxeno Olefinchemie GmbH, butis carried out using MTBE of standard industrial quality (fuel grade).Industrial MTBE usually contains C₈ olefins, for example those mentionedabove, 2-methoxybutane formed from linear butenes and methanol in thesynthesis of MTBE, TBA, methanol and possibly C₄- and C₅-hydrocarbons assecondary components.

Various processes for preparing isobutene by dissociation of MTBE areknown.

In EP 0 302 336, isobutene is obtained by dissociation of MTBE in acolumn which is a combination of a stirred vessel, a distillation columnand an extraction column. The acid catalyst is located in the bottomscircuit. MTBE is fed into the bottoms circuit. Part of the circulatingbottoms is taken off. The crude isobutene is extracted with water in thecolumn. The aqueous, methanol-containing extract is taken off above thebottom. The isobutene is obtained as overhead product. The bottoms takenoff are recirculated after discharge of a proportion thereof.

EP 0 068 785 claims a process in which the dissociation of MTBE iscarried out in a continuously operated stirred vessel. Here, the acidcatalyst is suspended in the starting material. Isobutene is separatedoff from the reaction mixture which distils off by means of a columnfrom which it is obtained as overhead product. Part of the bottomproduct is recirculated to the stirred vessel. The other part isseparated in a two-column system into an MTBE-containing stream which isrecirculated to the stirred vessel and methanol which is taken off as aside stream. The way in which secondary components in the MTBE used andby-products formed are separated off is not disclosed.

In DE 32 10 435, MTBE is dissociated in a reactive distillation column.An isobutene mixture containing small amounts of methanol and traces ofdiisobutene is obtained as overhead product. The bottom product from thereactive distillation column is separated in a distillation column intoan MTBE-containing stream which is recirculated to the reactivedistillation column and a bottom product comprising methanol. The way inwhich by-products are separated off is not indicated.

The documents EP 0 633 048, DE 100 20 943 and DE 101 11 549 claim thepreparation of isobutene by dissociation of MTBE in a reactivedistillation column. The separation of secondary components from thedissociation mixture is not disclosed.

DE 102 27 350 and DE 102 27 351 describe processes for preparingisobutene by dissociation of MTBE in the gas phase. In both processes,the dissociation product is condensed and extracted with water. Awater/methanol mixture is obtained as extract and is separated intomethanol and water by distillation. The raffinate comprises isobutene,unreacted MTBE and secondary components. This mixture is separated bydistillation into an overhead product which contains isobutene togetherwith small amounts of dimethyl ether and a bottom product comprisingMTBE and secondary components. No information is given about the use orwork-up of the MTBE mixture obtained.

U.S. Pat. No. 6,049,020 describes, inter alia, the preparation ofisobutene by dissociation of MTBE. Methanol is removed from the reactionproduct by extraction with water. The remaining raffinate is separatedby distillation into an overhead product containing isobutene and abottom product comprising unreacted MTBE and secondary components. Thework-up of the MTBE mixture is not described. In U.S. Pat. Nos.6,072,095 and 6,143,936, the work-up of the dissociation product iscarried out analogously. The MTBE mixture containing the secondarycomponents which is obtained is not worked up; it can be fed to a plantfor preparing MTBE.

In the dissociation of industrial MTBE which contains 2-methoxybutane,linear butenes are formed by dissociation of 2-methoxybutane. None ofthe abovementioned processes provides a solution to the separation oflinear butenes from isobutene or a limitation of the content of linearbutenes in the isobutene fraction.

In Catalysis Today 34 (1997), pages 447 to 455 (Case history: synthesisand decomposition of MTBE), R. Trotta and I. Miracca demonstrate thedependence of the content of linear butenes in isobutene which has beenproduced by dissociation of MTBE on the 2-methoxybutane content of thefeed MTBE and on the MTBE conversion. The relationship is shown in agraph which does not have a scale. The higher the 2-methoxybutanecontent of the feed MTBE and the higher the MTBE conversion, the higherthe content of linear n-butenes in the dissociation product. It is notpossible to prepare isobutene having a content of linear butenes of lessthan 1000 ppm by mass from a feed MTBE having a content of2-methoxybutane (MSBE) of over 1000 ppm by mass at quantitative MTBEconversion. In this case, the MTBE conversion has to be reduced. Thishas the consequence that a relatively large amount of a mixture whichcomprises mainly methanol and MTBE and has a higher concentration of2-methoxybutane than the feed MTBE is obtained after the isobutene hasbeen separated off. This mixture must not, because of its relativelyhigh 2-methoxybutane content, be fed into the dissociation reactor ifthe isobutene specification is to be adhered to. Even recirculation ofthis stream to an upstream MTBE plant which produces just the amount ofMTBE required for the dissociation is not possible, since2-methoxybutane would accumulate in the plant system. A purge streamassociated with MTBE losses is therefore unavoidable.

SUMMARY OF THE INVENTION

It was therefore an object of the present invention to provide a processfor the dissociation of MTBE, by means of which the dissociation of2-methoxybutane-containing MTBE can be carried out so that the isobuteneobtained contains less than 1000 ppm by mass of linear butenes, based onthe C₄-olefin fraction, and a purge stream for the secondary componentscan be kept small.

This and other objects have been achieved by the present invention thefirst embodiment of which includes a process for the dissociation ofmethyl tert-butyl ether (MTBE), comprising:

a) separating an MTBE-containing stream I by distillation into anMTBE-containing overhead stream II and a bottom stream III whichcomprises compounds having boiling points higher than that of MTBE; and

b) dissociation of the MTBE present in the overhead stream II over acatalyst to give a dissociation product IV;

wherein the stream I has a proportion of 2-methoxybutane (MSBE) ofgreater than 1000 ppm by mass, based on MTBE, and

wherein the separation by distillation in step a) and/or thedissociation in step b) is carried out so that the dissociation productIV has a concentration of less than 1000 ppm by mass of linear butenes,based on a C₄-olefin fraction.

BRIEF DESCRIPTION OF DRAWINGS

FIG. 1 shows a block diagram of an embodiment of a plant in which theprocess of the invention can be carried out.

FIG. 2 shows a block diagram of another embodiment of a plant in whichthe process of the invention can be carried out.

DETAILED DESCRIPTION OF THE INVENTION

It has now surprisingly been found that an MTBE dissociation processwhich comprises a step in which high boilers, in particular2-methoxybutane, are separated off by distillation and subsequently anMTBE dissociation step can be carried out so that the content of linearbutenes (which are obtained by dissociation of MSBE) in the dissociationproduct can be reduced to less than 1000 ppm by mass based on theisobutene.

This finding is particularly surprising because R. Trotta and I Miraccain Catalysis Today 34 (1997), page 453, state that 2-methoxybutane andMTBE have the same boiling point and can therefore not be separated fromone another.

The present invention accordingly provides a process for thedissociation of methyl tert-butyl ether (MTBE), which comprises at leastthe steps of separation by distillation of an MTBE-containing stream Iinto an MTBE-containing overhead stream II and a bottom stream III whichcomprises compounds having boiling points higher than that of MTBE anddissociation of the MTBE present in the overhead stream II over acatalyst to give a dissociation product IV, which is characterized inthat the stream I has a proportion of 2-methoxybutane (MSBE) of greaterthan 1000 ppm by mass based on MTBE and in that the separation bydistillation in step a) and/or the dissociation in step b) is carriedout so that the dissociation product IV obtained has a concentration ofless than 1000 ppm by mass of linear butenes, based on the C₄-olefinfraction.

The process of the present invention has the following advantages: theprocess is very flexible in respect of the various MTBE grades which canhave different concentrations of by-products such as diisobutene or2-methoxybutane. Adaptation of the distillation and/or dissociationconditions makes it possible in each case to produce an isobutene whichcontains less than 1000 ppm by mass of linear butenes, based on theC₄-olefin fraction. The mixture containing methanol, MTBE andby-products which remains after the isobutene has been separated offfrom the dissociation product can be recirculated in its entirety to anMTBE plant. This applies even when this MTBE plant produces only theamount of MTBE required for the dissociation. In the case of stand-aloneplants in which MTBE delivered from outside is used, it is possible torecirculate the remainder of the dissociation product which remainsafter isobutene and methanol have been separated off to the distillationcolumn.

The invention is described by way of example below without theinvention, whose range of protection is defined by the claims and thedescription, being restricted thereto. The claims themselves are alsopart of the disclosure content of the present invention. If ranges,general formulae or classes of compounds are mentioned below, thedisclosure encompasses not only the corresponding ranges or groups ofcompounds which are explicitly mentioned but also all subranges andsubsets of compounds which can be obtained by leaving out individualvalues (ranges) or compounds, without these having been explicitlymentioned for reasons of clarity.

The process of the invention for the dissociation of methyl tert-butylether (MTBE), which comprises at least the steps

a) separation by distillation of an MTBE-containing stream I into anMTBE-containing overhead stream II and a bottom stream III whichcomprises compounds having boiling points higher than that of MTBE and

b) dissociation of the MTBE present in the overhead stream II over acatalyst to give a dissociation product IV,

wherein

-   -   the stream I has a proportion of 2-methoxybutane (MSBE) of        greater than 1000 ppm by mass based on MTBE; and    -   the separation by distillation in step a) and/or the        dissociation in step b) is carried out so that the dissociation        product IV obtained has a concentration of less than 1000 ppm by        mass of linear butenes, based on the C₄-olefin fraction.        Process Step a)

The separation by distillation of the MTBE-containing stream I into anMTBE-containing overhead stream II and a bottom stream III comprisingcompounds having boiling points higher than that of MTBE in process stepa) is carried out in at least one column, preferably in precisely onedistillation column. In the column, the MTBE-containing stream I ispreferably fractionally distilled to separate off high boilers, inparticular 2-methoxybutane, from the MTBE. The distillation column usedpreferably has from 50 to 140 theoretical plates, more preferably from60 to 120 theoretical plates and most preferably from 80 to 110theoretical plates. The reflux ratio, which for the purposes of thepresent invention is defined as the mass flow of runback divided by themass flow of distillate, is, depending on the number of theoreticalplates realized, the composition of the MTBE used and the requiredpurity, preferably from 1 to 20, more preferably from 1 to 10. Theoperating pressure can be from 0.1 to 2.0 MPa_((abs)). The pressureincludes all values and subvalues therebetween, especially including0.2, 0.3, 0.4, 0.5, 0.6, 0.7, 0.8, 0.9, 1, 1.1, 1.2, 1.3, 1.4, 1.5, 1.6,1.7, 1.8, and 1.9 MPa_((abs)).

If the dissociation of the overhead product II is carried out in the gasphase at superatmospheric pressure in the dissociation reactor, it canbe advantageous to carry out the distillation at a higher pressure. Inthis case, the overhead condenser is preferably operated as a partialcondenser and the overhead product (II) is taken off in vapour form. Ifthe reaction pressure in the dissociation reactor is, for example, 0.7MPa_((abs)), the distillation pressure should preferably be at least0.75 MPa_((abs)). At operating pressures of greater than 1.3MPa_((abs)), ND steam can be generated by means of the heat ofcondensation and can be used for heating other columns of the process.Steam or heat transfer fluid can be used for heating the column,depending on the operating pressure selected. The bottom product III cancontain MTBE in addition to the high boilers such as 2-methoxybutane anddiisobutene. This mixture can be utilized thermally, be employed asstarting material for a synthesis gas plant or be used directly or afterhydrogenation as fuel component.

The content of 2-methoxybutane in the overhead stream II can be set todifferent values by variation of the distillation conditions. Dependingon the degree of dissociation of MTBE in the downstream reactor R,different 2-methoxybutane concentrations are permissible in the overheadstream II. If a very complete conversion of the MTBE present in thestream II is sought in the reactor R, the content of 2-methoxybutane instream II should not exceed 1000 ppm by mass in order to adhere to theisobutene specification of less than 1000 ppm by mass of linear butenes,since at a quantitative MTBE conversion the ratio of 2-methoxybutane toMTBE corresponds approximately to the ratio of linear butenes toisobutene in the dissociation product. On the other hand, if a lowerMTBE conversion is set in the reactor R, the 2-methoxybutaneconcentration in the stream II can be more than 1000 ppm by mass, sinceMTBE is dissociated more quickly than 2-methoxybutane.

The ratio of the dissociation rate of MTBE to that of 2-methoxybutanecan be dependent not only on the dissociation conditions but also on thecatalyst used. The 2-methoxybutane concentrations which are permissiblein the distillate II can be determined for the given catalyst andconversion by means of simple preliminary tests. For example, when thecatalysts described below, which formally comprise magnesium oxide,aluminum oxide and silicon oxide, are used, a 2-methoxybutaneconcentration of up to 2500 ppm by mass, in particular a concentrationof 2000 ppm by mass (in each case based on MTBE), is permissible at MTBEconversions of from 85 to 95% without the isobutene specification beingendangered.

In the present invention “conversion” refers to how much of a reactanthas reacted.

Process Step b)

The dissociation of the MTBE present in stream II into isobutene andmethanol can be carried out in the liquid phase or gas/liquid mixedphase or in the gas phase in the presence of acid catalysts. In theprocess of the invention, the MTBE dissociation is preferably carriedout in the gas phase. The MTBE dissociation is preferably carried out ata temperature in the range from 200 to 400° C., preferably from 230 to350° C. The temperature includes all values and subvalues therebetween,especially including 220, 240, 250, 260, 280, 300, 320, 340, 350, 360,380° C.

In the process of the invention, the dissociation of MTBE can be carriedout using all known acid catalysts which are suitable for thedissociation of MTBE. Acid catalysts which can be used are, for example,metal oxides, mixed metal oxides, in particular ones which containsilicon oxide or aluminum oxide, acids on metal oxide supports or metalsalts.

The dissociation in step b) is preferably carried out over a catalystwhich has an activity in respect of the dissociation of MTBE which is atleast 1%, preferably 5% and particularly preferably 10%, greater thanthe activity in respect of the dissociation of 2-methoxybutane. Theactivity of the catalyst can be determined in a simple manner byreacting a mixture of MSBE and MTBE under steady-state conditions overthe selected catalyst and subsequently analysing the resulting reactionproduct for unreacted MTBE and MSBE.

In the process of the invention for the dissociation of MTBE, preferablyfor the dissociation of MTBE in the gas phase, preference is given tousing catalysts which formally comprise magnesium oxide, aluminum oxideand silicon oxide. Such catalysts are described, for example, in U.S.Pat. No. 5,171,920 in Example 4 and in EP 0 589 557.

Particular preference is given to using catalysts which formallycomprise magnesium oxide, aluminum oxide and silicon dioxide and have aproportion of magnesium oxide from 0.5 to 20% by mass, preferably from 5to 15% by mass and particularly preferably from 10 to 15% by mass, aproportion of aluminum oxide of from 4 to 30% by mass, preferably from10 to 20% by mass, and a proportion of silicon dioxide of from 60 to 95%by mass, preferably from 70 to 90% by mass. The amount of magnesiumoxide includes all values and subvalues therebetween, especiallyincluding 1, 2, 4, 5, 6, 8, 10, 12, 14, 15, 16 and 18% by mass. Theamount of aluminum oxide includes all values and subvalues therebetween,especially including 5, 6, 8, 10, 12, 14, 15, 16, 18, 20, 22, 24, 25,26, and 28% by mass. The amount of silicon dioxide includes all valuesand subvalues therebetween, especially including 65, 70, 75, 80, 85, and90% by mass.

It can be advantageous for the catalyst to contain an alkali metal oxidein addition to the magnesium oxide. This alkali metal oxide can, forexample, be selected from among Na₂O and K₂O. The catalyst preferablycontains Na₂O as alkali metal oxide. The preferred catalyst preferablyhas a BET surface area (determined volumetrically by means of nitrogenin accordance with DIN ISO 9277) of from 200 to 450 m²/g, morepreferably from 200 to 350 m²/g. The BET includes all values andsubvalues therebetween, especially including 220, 240, 250, 260, 280,300, 320, 340, 350, 360, 380, 400, 420, 440 m²/g. If the catalyst usedaccording to the invention is applied as active composition to asupport, only the active composition has a BET surface area in the rangementioned. On the other hand, the material composed of catalyst andsupport can, depending on the nature of the support, have a BET surfacearea which deviates significantly from these values, in particular asmaller BET surface area.

The pore volume of the catalyst is preferably from 0.5 to 1.3 ml/g, morepreferably from 0.65 to 1.1 ml/g. The pore volume is preferablydetermined by the cyclohexane method. In this method, the sample to betested is firstly dried to constant weight at 110° C. About 50 ml of thesample weighted to within 0.01 g are subsequently introduced into animpregnation tube which has been cleaned and dried to constant weightand has an outlet opening provided with a ground glass stopcock on theunderside. The outlet opening is covered with a small polyethylene plateso as to prevent blockage of the outlet opening by the sample. Aftercharging of the impregnation tube with the sample, the tube is carefullyclosed so as to be airtight. The impregnation tube is subsequentlyconnected to a water pump, the ground glass stopcock is opened and avacuum of 20 mbar is set in the impregnation tube by means of the waterpump. The vacuum can be checked on a vacuum gauge connected in parallel.After 20 minutes, the ground glass stopcock is closed and the evacuatedimpregnation tube is subsequently connected to a cyclohexane reservoirwhich has been charged with a precisely measured volume of cyclohexaneso that cyclohexane is sucked into the impregnation tube from thereservoir on opening the ground glass stopcock. The ground glassstopcock remains open until the entire sample is flooded withcyclohexane. The ground glass stopcock is subsequently closed again.After 15 minutes, air is carefully admitted into the impregnation tubeand the cyclohexane which has not been absorbed is drained into thereservoir. Cyclohexane adhering to the inside of the impregnation tubeor to the outlet opening or the connection with the cyclohexanereservoir can be conveyed into the reservoir by means of a singlecareful pressure pulse from a rubber bulb via the air admission line.The volume of the cyclohexane present in the reservoir is noted. Thepore volume is given by the volume of cyclohexane absorbed, which isdetermined from the volume of cyclohexane in the reservoir before themeasurement minus the volume of cyclohexane in the reservoir after themeasurement, divided by the mass of the sample examined.

The mean pore diameter (preferably determined by a method based on DIN66133) of the catalyst is preferably from 5 to 20 nm, more preferablyfrom 8 to 15 nm. Particular preference is given to at least 50%,preferably more than 70%, of the total pore volume (sum of the porevolume of pores having a pore diameter of greater than or equal to 3.5nm determined by mercury porosimetry in accordance with DIN 66133) ofthe catalyst being made up by pores having a diameter of from 3.5 to 50nm (mesopores).

The process of the invention is preferably carried out using catalystswhich have a mean particle size (determined by sieve analysis) of from10 μm to 10 mm, preferably from 0.5 mm to 10 mm, particularly preferablya mean particle size of from 1 to 5 mm. Preference is given to usingsolid catalysts which have a mean particle size d₅₀ of from 2 to 4 mm,in particular from 3 to 4 mm.

The catalyst can be used as shaped bodies in the process of theinvention. The shaped bodies can have any shape. The catalyst ispreferably used as shaped bodies in the form of spheres, extrudates orpellets. The shaped bodies preferably have the abovementioned meanparticle sizes.

The catalyst can also be applied to a support, e.g. a metal, plastic orceramic support, preferably a support which is inert in respect of thereaction for which the catalyst is to be used. In particular, thecatalyst used according to the invention can be applied to a metalsupport, e.g. a metal plate or a metal mesh. Such supports provided withthe catalyst used according to the invention can, for example, be usedas internals in reactors or reactive distillation columns. The supportscan also be metal, glass or ceramic spheres or spheres of inorganicoxides. If the catalyst used according to the invention is applied to aninert support, the mass and composition of the inert support are nottaken into account in determining the composition of the catalyst.

The particularly preferred catalysts which formally comprise magnesiumoxide, aluminum oxide (Al₂O₃) and silicon dioxide can, for example, beproduced by a process comprising the steps

A1) treatment of an aluminosilicate with an acidic, aqueous magnesiumsalt solution and

B1) calcination of the aluminosilicate treated with aqueous magnesiumsalt solution.

For the purposes of the present invention, aluminosilicates arecompounds which are formally composed essentially of proportions ofaluminum oxide (Al₂O₃) and silicon dioxide (SiO₂). However, thealuminosilicates can also contain small proportions of alkali metaloxides or alkaline earth metal oxides. Zeolites such as zeolites A, X,Y, USY or ZSM-5 or amorphous zeolites (for example MCM 41 from MobilOil) can also be used as aluminosilicates in the process. Thealuminosilicates used in the process can be amorphous or crystalline.Suitable commercial aluminosilicates which can be used as startingmaterials in the process of the invention are, for example,aluminosilicates which have been prepared by precipitation, gelation orpyrolysis. The process is preferably carried out using aluminosilicateswhich comprise from 5 to 40% by mass, preferably from 10 to 35% by mass,of aluminum oxide and from 60 to 95% by mass, preferably from 65 to 90%by mass, of silicon dioxide (based on the dry mass; treatment: ignitionat 850° C. for 1 hour). The amount of aluminum oxide includes all valuesand subvalues therebetween, especially including 5, 6, 8, 10, 12, 14,15, 16, 18, 20, 22, 24, 25, 26, 28, 30, 32, 34, 35, 36, and 38% by mass.The amount of silicon dioxide includes all values and subvaluestherebetween, especially including 65, 70, 75, 80, 85, and 90% by mass.

The composition of the aluminosilicates used or the catalysts obtainedcan, for example, be determined by classical analysis, fusion with Boraxand XRF (X-ray fluorescence), energy-dispersive X-ray analysis, flamespectroscopy (Al and Mg, not Si), wet digestion and subsequent ICP-OES(Optical Emission Spectrometry with Inductively Coupled High-FrequencyPlasma) or atomic absorption spectroscopy. A particularly preferredaluminosilicate which can be used in the process has a formal proportionof Al₂O₃ of 13% by mass and a proportion of silicon dioxide of 76% bymass. Such an aluminosilicate is marketed by Grace Davison under thetrade name Davicat O 701.

The aluminosilicate can be used in a variety of forms in the process.Thus, the aluminosilicate can be used in the form of shaped bodies suchas tablets, pellets, granules, rods or extrudates. However, thealuminosilicate can also be used as aluminosilicate powder. As startingmaterial, it is possible to use powders having various mean particlesizes and various particle size distributions. In the production ofshaped bodies, preference is given to using an aluminosilicate powder inwhich 95% of the particles have a mean particle size of from 5 to 100μm, preferably from 10 to 30 μm and particularly preferably from 20 to30 μm. The particle size can be determined, for example, by laser lightscattering using a particle analyser from Malvern, e.g. the Mastersizer2000.

The aqueous magnesium salt solution is produced using magnesiumcompounds which are water-soluble or are converted into water-solublecompounds by addition of an acid. The nitrates are preferably used assalts. Preference is given to using magnesium salt solutions whichcontain the salts of strong mineral acids, for example magnesium nitratehexahydrate or magnesium sulphate heptahydrate, as magnesium salts. Theacidic aqueous alkali metal and/or alkaline earth metal salt solutionused preferably has a pH of less than 6, more preferably from <6 to 3and particularly preferably from 5.5 to 3.5. The pH can be determined,for example, by means of a glass electrode or indicator paper. If thesalt solution has a pH which is greater than or equal to 6, the pH canbe adjusted by addition of an acid, preferably the acid whose alkalimetal and/or alkaline earth metal salt is present in the solution. Whenthe alkali metal and/or alkaline earth metal salt solution contains thenitrates as salts, nitric acid is preferably used as acid. The magnesiumcontent of the magnesium salt solution used is preferably from 0.1 to 3mol/l, preferably from 0.5 to 2.5 mol/l.

The treatment in step A1) can be carried out in various ways which aresuitable for bringing the aluminosilicate into contact with themagnesium salt solution. Possible treatment methods are, for example,impregnation, steeping, spraying or irrigation of the aluminosilicatewith the magnesium salt solution. It can be advantageous for thetreatment of the aluminosilicate to be carried out so that the magnesiumsalt solution can act on the aluminosilicate for at least from 0.1 to 5hours, preferably from 0.5 to 2 hours. Such a contact time can beparticularly advantageous when the treatment is carried out by simplesteeping.

In a preferred embodiment of step A1) of the process, the treatment ofaluminosilicate, in particular shaped aluminosilicate bodies, with themagnesium salt solution can be carried out, for example, by vacuumimpregnation in a vacuum impregnation unit suitable for this purpose. Inthis type of treatment, the aluminosilicate is firstly evacuated in thevacuum impregnation unit. The magnesium salt solution is subsequentlysucked in up to above the upper surface of the bed of the support, sothat all the aluminosilicate is covered with the solution. After acontact time which is preferably from 0.1 to 10 hours, more preferablyfrom 0.5 to 2 hours, the solution which has not been taken up by thesupport is drained.

In a further preferred embodiment of step A1) of the process, thetreatment of aluminosilicate, in particular shaped aluminosilicatebodies, with the alkali metal and/or alkaline earth metal salt solution,can be carried out, for example, by spraying or irrigation of thealuminosilicate. The spraying or irrigation of the aluminosilicate withthe magnesium salt solution is preferably carried out by spraying orpouring the solution onto the aluminosilicate rotating in a drum. Thetreatment can be carried out in one action, i.e. the total amount ofmagnesium salt solution is added to the aluminosilicate at the beginningin one step. However, the salt solution can also be introduced in smallportions by spraying or irrigation, the time of addition preferablybeing from 0.1 to 10 hours and more preferably from 1 to 3 hours. Theamount of salt solution is preferably measured so that all of thesolution is taken up by the aluminosilicate. Steeping in particular butalso spraying or irrigation can be carried out in conventionalindustrial apparatuses, for example cone mixers or high-intensity mixersas are, for example, marketed by Eirich.

The treatment of the aluminosilicate with the magnesium salt solution instep A1) can be carried out in one step or in a plurality of substeps.In particular, it is possible to carry out the treatment in two or moresubsteps. In each of the individual substeps, it is possible to use thesame magnesium salt solution in each case or else to use a magnesiumsalt solution having a different concentration in each substep. Forexample, it is possible to initially add only part of the magnesium saltsolution to the aluminosilicate and, if desired after intermediatedrying, to add the remainder of the magnesium salt solution to be usedat the same temperature or a different temperature. It is not onlypossible for step A1) to be carried out in two or more substeps. It islikewise possible for the process to have a plurality of steps A1). Inthis case too, magnesium salt solutions having the same concentration ordifferent concentrations can be used in the various steps A1).

The treatment in step A1) is preferably carried out at a temperature offrom 10 to 120° C., more preferably from 10 to 90° C., particularlypreferably from 15 to 60° C. and very particularly preferably at atemperature of from 20 to 40° C.

It can be advantageous for one or more additives to be added to or mixedinto the aluminosilicate or the magnesium salt solution in step A1).Such additives can be, for example, binders, lubricants or shaping aids.A suitable binder can be, for example, boehmite or pseudoboehmite, as ismarketed, for example, under the trade name Disperal (a boehmite havingan Al₂O₃ content of about 77% by mass (balance water and traces ofimpurities) by Sasol Deutschland GmbH. If boehmite, in particularDisperal, is added as binder, it is preferably added as a gel which canbe obtained by, for example, stirring 197 parts by mass of Disperal into803 parts by mass of a 1.28% strength by mass aqueous nitric acid,stirring vigorously at 60° C. for 3 hours, cooling to room temperatureand replacing any water which has evaporated. As shaping aids, it ispossible to use, for example, silicas, in particular pyrogenic silicasas are marketed, for example, by Degussa AG under the trade nameAerosil, bentonites, clays, kaolin, kaolinite, ball clay and othermaterials with which those skilled in the art are familiar for thispurpose. As lubricant, whose use can be advantageous for improvedtableting, it is possible to add, for example, graphite.

The addition of one or more of the abovementioned additives in step A1)can be carried out in various ways. In particular, the addition can becarried out during the treatment of the aluminosilicate with themagnesium salt solution. For example, aluminosilicate, additive andmagnesium salt solution can be introduced into an industrial apparatusand subsequently mixed intimately. Another possibility is firstly to mixthe aluminosilicate with the additive and subsequently add the magnesiumsalt solution. In a further variant, additive and magnesium saltsolution can be added simultaneously to the aluminosilicate. Theaddition can in each case be carried out in one addition action, inportions or by spraying. The addition time is preferably less than 5hours, more preferably less than 3 hours. It can be advantageous to mixthe mixture for a further 0.1-10 hours, preferably 0.5-3 hours.

The process for producing the catalyst which is preferably used has atleast one process step B1) in which the aluminosilicate which has beentreated with alkali metal and/or alkaline earth metal salt solution iscalcined. The calcination is preferably carried out in a gas stream, forexample a gas stream containing, for example, air, nitrogen, carbondioxide and/or one or more noble gases or consists of one or more ofthese components. The calcination is preferably carried out using air asgas stream.

The calcination in process step B1) is preferably carried out at atemperature of from 200 to 1000° C., more preferably from 300 to 800° C.The calcination is preferably carried out for a time of from 0.1 to 10hours, more preferably from 1 to 5 hours. The calcination isparticularly preferably carried out at a temperature of from 200 to1000° C., preferably from 300 to 800° C., for from 0.1 to 10 hours,preferably from 1 to 5 hours.

Industrial calcination can preferably be carried out in a shaft furnace.However, the calcination can also be carried out in other knownindustrial apparatuses, for example fluidized-bed calciners, rotary tubefurnaces or tray furnaces.

It can be advantageous for a step C1) in which the aluminosilicate whichhas been treated with magnesium salt solution is dried to be carried outbetween steps A1) and B1). Drying in step C1) can be carried out at atemperature of from 100 to 140° C. Drying is preferably carried out in agas stream. Drying can, for example, be carried out in a gas streamwhich contains, for example, air, nitrogen, carbon dioxide and/or one ormore noble gases or consists of one or more of these components. Theintermediate step of drying after treatment with alkali metal and/oralkaline earth metal salt solution and before calcination makes itpossible to ensure that large amounts of water vapour are not liberatedduring calcination. In addition, drying can prevent water fromevaporating spontaneously during calcination and destroying the shape ofthe catalyst.

Depending on the desired shape of the catalyst, it can be advantageousto appropriately adapt the production process by means of additionalprocess steps. If, for example, a pulverulent catalyst is to be producedby the process, the aluminosilicate can be used in the form ofaluminosilicate powder and, for example, be treated with the magnesiumsalt solution (e.g. by impregnation) in a cone mixer, optionally driedand subsequently calcined. However, a pulverulent catalyst can also beproduced by processing a shaped catalyst body by milling and sieving togive a pulverulent catalyst.

The shaped catalyst bodies can, for example, be in the form ofextrudates, spheres, pellets or tablets. To obtain the shaped catalyst(shaped catalyst body), further process steps such as shaping, millingor sieving can be carried out in addition to the process steps oftreatment, drying and calcination, depending on the respective shapingvariant. Shaping aids can be introduced into the process at variouspoints. The shaped catalyst bodies can be produced in various ways:

In a first variant, shaped catalyst bodies, in particular shapedcatalyst bodies to be used according to the invention, can be obtainedby treating shaped aluminosilicate bodies with an acidic aqueousmagnesium salt solution, optionally drying them and subsequentlycalcining them.

In a second embodiment, a catalyst body can be obtained by firstlytreating an aluminosilicate powder with an acidic aqueous magnesium saltsolution, then drying it if desired and subsequently calcining it andthen processing the resulting catalyst powder by industrially customarymethods such as compacting, extrusion, pelletization, tableting,granulation or coating to produce shaped catalyst bodies. Additivesrequired for shaping, e.g. binders or further auxiliaries, can be addedat various points in the production process, e.g. in process step A1).In the production of a shaped body from an aluminosilicate powder asstarting material, it is possible to start out from powders havingvarious mean particle sizes and various particle size distributions. Inthe production of shaped bodies, preference is given to using analuminosilicate powder in which 95% of the particles have a particlesize of from 5 to 100 μm, preferably from 10 to 30 μm and particularlypreferably from 20 to 30 μm (determined by laser light scattering, seeabove).

In a third embodiment of the process, pellets of the catalyst can beobtained by treating an aluminosilicate powder with an acidic aqueousmagnesium salt solution in process step A1), optionally drying it(process step C1)) and subsequently calcining it in process step B1) andthen pelletizing the resulting catalyst powder, e.g. in an Eirich mixer,with addition of binders and drying the resulting pellets in a furtherprocess step C1) and subsequently calcining them in a further processstep B1).

In a fourth embodiment of the production process, pellets of thecatalyst can be obtained by mixing an aluminosilicate powder, binder andacidic aqueous magnesium salt solution in process step A1) andpelletizing the treated aluminosilicate powder, e.g. in an Eirich mixer,and drying the resulting moist pellets in process step C1) andsubsequently calcining them in a gas stream in process step B1).

In a fifth embodiment of the production process, tablets of the catalystcan be obtained by mixing an aluminosilicate powder, binder, optionallylubricant and acidic aqueous magnesium salt solution in process step A1)and pelletizing the treated aluminosilicate powder, e.g. in an Eirichmixer, to form micropellets having a mean diameter of preferably from0.5 to 10 mm, more preferably from 1 to 5 mm and particularly preferablyfrom 1 to 3 mm (determination of the particle size can, for example, becarried out by sieve analysis) and drying the resulting moist pellets inprocess step C1) and subsequently, if desired, calcining them in a gasstream in process step B1). The pellets obtained can then, if this hasnot yet occurred in process step A1), be mixed with a lubricant such asgraphite and subsequently be tableted on a commercial tableting press,e.g. a rotary press. The tablets can then, if a process step B1) has notyet been carried out, be calcined in process step B1).

In a sixth embodiment of the production process, tablets of the catalystcan be obtained by milling preshaped shaped catalyst bodies, as can beobtained, for example, as pellets in embodiment three and four, andsieving the granules/powder obtained so as to give a tabletable granularcatalyst material and mixing lubricant into this granular material. Thegranular material which has been prepared in this way can subsequentlybe tableted. The tablets can then, if a process step B1) has not yetbeen carried out, be calcined in process step B1). The addition of alubricant can be omitted if a lubricant has already been added in theproduction of the pellets, e.g. in process step A1).

In a seventh embodiment of the process of the invention,materials/supports coated with the catalyst can be produced. In thisembodiment, a catalyst powder will firstly be produced by treating analuminosilicate powder with an acidic aqueous magnesium salt solution inprocess step A1), optionally drying it (process step C1)) and optionallycalcining it (process step B1)). The catalyst powder obtained in thisway is subsequently suspended in a suspension medium such as water oralcohol, if desired with addition of a binder to the suspension. Thesuspension produced in this way can then be applied to any desiredmaterial. After application of the catalyst suspension, the material isoptionally dried (process step C1)) and subsequently calcined (processstep B1)). Materials/supports coated with the preferred catalyst can beprovided in this way. Such materials/supports can be, for example, metalplates or meshes as can be used as internals in reactors or columns, inparticular reactive distillation columns, or else metal, glass orceramic spheres or spheres composed of inorganic oxides.

In an eighth embodiment of the production process, extrudates of thecatalyst, in particular the catalyst to be used according to theinvention, can be obtained by mixing an aluminosilicate powder, acidicaqueous alkali metal and/or alkaline earth metal salt solution, bindersuch as Disperal and further shaping aids customary for extrusion, forexample clays such as bentonite or attapulgite, in process step A1) in akneader or Eirich mixer and extruding the mixture in an extruder to giveextrudates having a mean diameter of preferably from 0.5 to 10 mm, morepreferably from 1 to 5 mm and particularly preferably from 1 to 3 mm,and optionally drying the resulting moist extrudates in process step C1)and subsequently calcining them in a gas stream in process step B1).

The reaction pressure in step b) is preferably from 0.1 to 10 MPa(abs),preferably from 0.5 to 0.8 MPa(abs). The dissociation is preferablycarried out at a WHSV (weight hourly space velocity) of from 0.1 to 5h⁻¹, preferably from 1 to 3 h⁻¹ (kg of MTBE per kg of catalyst perhour). The MTBE conversion in a single pass is preferably from 70 to98%, more preferably from 90 to 95%. The dissociation can be carried outin customary reactors, e.g. in a tube reactor, shell-and-tube reactor,shaft furnace or fluidized-bed reactor or a combination thereof.

The dissociation in the gas phase is preferably carried out in a reactorwhich is provided with a heating jacket and is heated by means of aliquid heat transfer medium, the dissociation being carried out so thatthe temperature drop in the catalyst zone/reaction zone at any desiredpoint relative to the inlet temperature is less than 50° C., preferablyless than 40° C. and particularly preferably from 1 to 30° C., so thatthe reaction mixture in the reactor and the heat transfer medium in thejacket flow in cocurrent through the reactor and so that the differencein temperature of the heat transfer medium between inflow point to thereactor and exit from the reactor is less than 40° C. The maximumtemperature drop can be set by means of numerous parameters, e.g. bymeans of the temperature of the heat transfer medium used for heatingand via the rate at which the heat transfer medium flows through thejacket.

The inlet temperature of the gaseous starting material, in particular inthis preferred embodiment of process step b) according to the invention,is preferably above 200° C., more preferably above 230° C. andparticularly preferably above 250° C. The inlet temperature of thestarting material can be set in a heater located upstream of thereactor. When fresh catalyst is used in the MTBE dissociation, inparticular when fresh magnesium oxide/aluminum oxide/silicon oxidecatalyst is used, the inlet temperature is preferably in the range from250 to 270° C. It can be advantageous to increase the inlet temperatureup to 400° C. during operation as the catalyst becomes increasinglydeactivated in order to keep the conversion constant. When theconversion can no longer be kept constant after reaching 400° C., it canbe advantageous to replace all or part of the catalyst.

The reactor is, especially in the case of this preferred embodiment ofprocess step b) according to the invention, preferably operated at aspace velocity (weight hourly space velocity (WHSV) in kilogram ofstarting material per kilogram of catalyst per hour) of from 0.1 to 5h⁻¹, in particular from 1 to 3 h⁻¹, in a single pass.

The reactor can, especially in this preferred embodiment of process stepb) according to the invention, be arranged in any direction in space. Ifthe reactor has reaction tubes, these can likewise point in anydirection in space. However, the reactor is preferably erected so thatthe reactor or the reaction tubes are vertical. In the case of avertical reactor, the heat transfer medium is preferably fed in at thehighest point or in the vicinity of the highest point of the jacket andtaken off at the lowest point or in the vicinity of the lowest point ofthe reactor, or vice versa. The reaction mixture in the reaction zoneand the heat transfer medium in the jacket preferably flow through thereactor in the same direction. The heat transfer medium and the reactionmixture particularly preferably flow through the jacket of the reactorand the reaction zone of the reactor, respectively, from the topdownwards.

To achieve more uniform heating of the reaction zone, it can beadvantageous to feed the heat transfer medium into the reactor not onlyat one point but at a plurality of points at about the same height. Toavoid a relatively large temperature drop in the middle tubes comparedto tubes around the outside when a shell-and-tube reactor is used, itcan be advantageous to provide nozzles which favour transport of theheat transfer medium to the middle tubes in the inlet or inlets for theheat transfer medium. In this way, temperature fluctuations over thecross section of the bundle of tubes can be avoided.

The heat transfer medium can leave the reactor at one or more point(s).If the heat transfer medium flows through the reactor from the topdownwards, it should be ensured by constructional measures that thereaction zones, e.g. the reaction tubes, are completely surrounded byflowing heat transfer medium.

The heat transfer medium can be brought to the desired temperatureoutside the reactor by means of direct or indirect heating and pumpedthrough the reactor.

As heat transfer medium, it is possible to use salt melts, water or heattransfer fluids. Use of heat transfer fluids is advantageous for thetemperature range from 200 to 400° C., since heating circuits using themrequire a lower capital investment compared to other engineeringsolutions. Heat transfer fluids which can be used are, for example,those which are marketed under the trade names Marlotherm (e.g.Marlotherm SH from Sasol Olefins & Surfactants GmbH), Diphyl (fromBayer), Dowtherm (from Dow) or Therminol (from Therminol). Thesesynthetic heat transfer fluids are based essentially on thermally stablecyclic hydrocarbons.

The heat transfer medium is preferably fed into the heating jacket ofthe reactor at a temperature which is from 10 to 40° C., preferably from10 to 30° C., higher than the temperature of the starting materialflowing into the reactor. The difference in temperature of the liquidheat transfer medium over the reactor, i.e. between the inlettemperature of the heat transfer medium on entering the heating jacketand the outlet temperature of the heat transfer medium on leaving theheating jacket, is preferably less than 40° C., more preferably lessthan 30° C. and particularly preferably from 10 to 25° C. Thetemperature difference can be set via the mass flow of heat transfermedium per unit time (kilogram per hour) through the heating jacket.

The preferred embodiment of process step b) according to the inventioncan be carried out in all suitable reactors which are provided with aheating jacket and can be heated by means of a liquid heat transfermedium. Such reactors have a reaction zone in which the catalyst ispresent (catalyst zone) which is separated physically from a heatingjacket through which the heat transfer medium flows. The process of theinvention is preferably carried out in a plate reactor, in a tubereactor, in a plurality of tube reactors or plate reactors connected inparallel or in a shell-and-tube reactor. The process of the invention ispreferably carried out in a shell-and-tube reactor.

It may be pointed out that the hollow bodies in which the catalyst ispresent do not have to be tubes in the normal sense of the word. Thehollow bodies can also have noncircular cross sections. They can, forexample, be elliptical or triangular.

The materials for construction of the reactor, in particular thematerial which separates the reaction zone from the heating jacket,preferably have a high thermal conductivity coefficient (greater than 40W/(m° K)). Preference is given to using iron or an iron alloy, e.g.steel, as material having a high thermal conductivity coefficient.

If the process of the invention is carried out in a shell-and-tubereactor, the individual tubes preferably have a length of from 1 to 15m, more preferably from 3 to 9 m and particularly preferably from 5 to 9m. The individual tubes in a shell-and-tube reactor used in the processof the invention preferably have an internal diameter of from 10 to 60mm, more preferably from 20 to 40 mm and particularly preferably from 24to 35 mm. It can be advantageous for the individual tubes of theshell-and-tube reactor used in the process of the invention to have athickness of the tube wall of from 1 to 4 mm, preferably from 1.5 to 3mm.

In a shell-and-tube reactor used in the preferred embodiment of processstep b) according to the invention, the tubes are preferably arranged inparallel. The tubes are preferably arranged uniformly. The arrangementof the tubes can be, for example, square, triangular or diamond-shaped.Particular preference is given to an arrangement in which the virtuallyconnected middle points of three mutually adjacent tubes form anequilateral triangle, i.e. the tubes have the same spacing. The processof the invention is preferably carried out in a shell-and-tube reactorin which the tubes have a spacing of from 3 to 15 mm, particularlypreferably from 4 to 7 mm.

The dissociation in step b) is preferably carried out under conditionsunder which the conversion of MTBE is greater than the conversion of2-methoxybutane. In this simple fashion, it can be ensured that thedissociation product has a concentration of less than 1000 ppm by massof linear butenes based on the C₄-olefin fraction, even when theoverhead stream II comprises a proportion of 2-methoxybutane (MSBE) ofgreater than 1000 ppm by mass based on MTBE. Depending on the proportionof MSBE present in the overhead stream II, it may be necessary to setconditions under which the conversion of MSBE is significantly, i.e.,for example, at least 50%, lower than the conversion of MTBE. Theseconditions can be determined by means of simple preliminary tests.

The main reaction in process step b) of the process of the invention isthe dissociation of MTBE into isobutene and methanol. When no recyclestream from the dissociation is admixed with the feed MTBE, thedissociation product IV, depending on the MTBE conversion set,preferably has a residual MTBE content of from 2 to 29% by mass. Themethanol content is preferably from 25 to 35% by mass, the isobutenecontent is preferably from 44 to 61% by mass.

When unreacted MTBE is separated from the dissociation product in adistillation following on from the dissociation and is recirculated tothe dissociation, the compositions change in accordance with themethanol content of the recirculated MTBE. In the dissociation, theformation of diisobutene from isobutene and the reaction of methanol toform dimethyl ether can occur as secondary reactions. 2-methoxybutanepresent in the starting material can also be partly dissociated to formlinear butenes and tert-butyl alcohol (TBA) present can be dissociatedinto isobutene and water. Diisobutene, dimethyl ether, linear butenesand water, inter alia, can therefore be present in the dissociationproduct IV as further components formed by reaction in process step b).

Process Step c)

To work up the dissociation product mixture further, it may beadvantageous for the dissociation product IV to be separated in afurther distillation step c) into an isobutene-containing overheadstream V and a bottom stream VI comprising unreacted MTBE. Theseparation of the dissociation product IV into an isobutene-containingoverhead stream V and a bottom stream VI comprising unreacted MTBE bydistillation in process step c) is carried out in at least one column,preferably precisely one distillation column.

A distillation column which is preferably used in process step c)preferably has from 20 to 55 theoretical plates, more preferably from 25to 45 theoretical plates and particularly preferably from 30 to 40theoretical plates. The reflux ratio is, depending on the number oftheoretical plates realized, the composition of the output from thereactor and the required purities of distillate and bottom product,preferably less than 5, more preferably less than 1. The operatingpressure of the column K2 can preferably be set to from 0.1 to 2.0MPa(abs). The pressure includes all values and subvalues therebetween,especially including 0.2, 0.3, 0.4, 0.5, 0.6, 0.7, 0.8, 0.9, 1, 1.1,1.2, 1.3, 1.4, 1.5, 1.6, 1.7, 1.8, and 1.9 MPa_((abs)).

To save a compressor, it can be advantageous to operate the column at apressure which is lower than the pressure at which the dissociationreactor R in process step b) is operated. To be able to condenseisobutene by means of cooling water, a pressure of about 0.5 MPa(abs) isnecessary. If the dissociation in process step b) is carried out at, forexample, a pressure of 0.65 MPa(abs), it can be advantageous for thedistillation column of process step c) to be operated at a pressure offrom 0.55 to 0.6 MPa(abs). Heating of the vaporizer can be effectedusing, for example, 0.4 MPa steam. The bottom product VI preferablycontains unreacted MTBE, methanol and possibly by-products such asdiisobutene and 2-methoxybutane. The overhead product is preferablyisobutene having a purity of greater than 95% by mass, based on thetotal overhead product.

Process step c) can optionally be carried out in at least one columnconfigured as a reactive distillation column. This embodiment of theprocess of the invention has the advantage that the MTBE conversion inthe overall process can be increased by part of the MTBE which has notbeen reacted in process step b) being dissociated into isobutene andmethanol in the reaction part of the reactive distillation column ofprocess step c).

As catalysts in the reaction part of the reactive distillation column,it is possible to use all catalysts which are suitable for thedissociation of MTBE. Preference is given to using acid catalysts ascatalysts. A particularly preferred group of acid catalysts for use inthe reaction part of the reactive distillation column are solid, acidicion-exchange resins, in particular ones having sulphonic acid groups.Suitable acidic ion-exchange resins are, for example, those which areprepared by sulphonation of phenol/aldehyde condensates or ofcooligomers of aromatic vinyl compounds. Examples of aromatic vinylcompounds for preparing the cooligomers are: styrene, vinyltoluene,vinylnaphthalene, vinylethylbenzene, methylstyrene, vinylchlorobenzene,vinylxylene and divinylbenzene. In particular, the cooligomers formed byreaction of styrene with divinylbenzene are used as precursor forpreparing ion-exchange resins having sulphonic acid groups. The resinscan be produced in gel, macroporous or sponge form. The properties ofthese resins, in particular the specific surface area, porosity,stability, swelling or shrinkage and exchange capacity, can be varied bymeans of the production process.

The ion-exchange resins can be used in their H form in the reactionsection of the reactive distillation column. Strong acid resins of thestyrene-divinylbenzene type are, inter alia, sold under the followingtrade names: Duolite C20, Duolite C26, Amberlyst 15, Amberlyst 35,Amberlyst 46, Amberlite IR-120, Amberlite 200, Dowex 50, Lewatit SPC118, Lewatit SPC 108, K2611, K2621, OC 1501.

The pore volume of the ion-exchange resins used is preferably from 0.3to 0.9 ml/g, in particular from 0.5 to 0.9 ml/g. The particle size ofthe resin is preferably from 0.3 mm to 1.5 mm, more preferably from 0.5mm to 1.0 mm. The particle size distribution can be selected so as to berelatively narrow or relatively broad. Thus, for example, it is possibleto use ion-exchange resins having a very uniform particle size(monodisperse resins). The capacity of the ion exchanger, based on theform as supplied, is preferably from 0.7 to 2.0 eq/l, in particular from1.1 to 2.0 eq/l.

In the reaction part of a column which has optionally been configured asa reactive distillation column in process step c), the catalyst caneither be integrated into the packing, for example into KataMax® (asdescribed in EP 0 428 265) or KataPak® (as described in EP 0 396 650 orDE 298 07 007.3 U1) packings, or be polymerized onto shaped bodies (asdescribed in U.S. Pat. No. 5,244,929).

The reactive distillation column preferably has a region of purelydistillative separation above the catalyst packing. The zone above thecatalyst packing preferably has from 5 to 25, in particular from 5 to15, theoretical plates. The separation zone below the catalystpreferably encompasses from 5 to 35, more preferably from 5 to 25,theoretical plates. The feed to the reactive distillation column can beintroduced above or below, preferably above, the catalyst zone.

The conversion of the MTBE into isobutene and methanol in the reactivedistillation is preferably carried out in a temperature range from 60 to140° C., more preferably from 80 to 130° C., particularly preferablyfrom 90 to 110° C. (temperature in the region of the column in which thecatalyst is present; the temperature at the bottom can be significantlyhigher).

As regards the operating pressure of the reactive distillation column,it is in principle possible to choose operating conditions similar tothose for the above-described embodiment as pure distillation column.Thus, preference is given to setting an operating pressure in thereactive distillation column of from 0.1 to 1.2 MPa(abs). To save acompressor, it can be advantageous to operate the column at a pressurelower than the pressure at which the dissociation reactor R in processstep b) is operated. To be able to condense isobutene by means ofcooling water, a pressure of about 0.5 MPa(abs) is necessary. If thedissociation in process step b) is, for example, carried out at apressure of 0.65 MPa(abs), it can be advantageous for the distillationcolumn of process step c) to be operated at a pressure of from 0.55 to0.6 MPa(abs). The vaporizer can be heated using, for example, steam.

The hydraulic loading in the catalytic packing of the column ispreferably from 10% to 110%, more preferably from 20% to 70%, of itsflooding point loading. For the purposes of the present invention, thehydraulic loading of a distillation column is the uniform hydrodynamicloading of the column cross section by the ascending vapour stream andthe descending liquid stream. The upper loading limit is the maximumloading by vapour and liquid runback above which the separating actiondecreases as a result of entrainment or backing-up of the liquid runbackby the ascending vapour stream. The lower loading limit is the minimumloading below which the separating action decreases or breaks down as aresult of irregular flow or empty running of the column, e.g. the trays.(Vauck/Müller, “Grundoperationen chemischer Verfahrenstechnik”, p. 626,VEB Deutscher Verlag für Grundstoffindustrie.)

When the column in process step c) is configured as a reactivedistillation column, the bottom product VI obtained preferably likewisecontains unreacted MTBE and methanol and possibly by-products such asdiisobutene and 2-methoxybutane. The overhead product preferablycomprises isobutene having a purity of greater than 95% by mass.

The bottom product VI obtained in process step c) contains the MTBEwhich has not been reacted in process step b), and the major part of themethanol formed in the dissociation of the MTBE. By-products, forexample diisobutene and/or 2-methoxybutane may also be present in thebottom product. There are various possibilities for using or working upthis stream VI. If the MTBE dissociation plant is associated with aplant for preparing MTBE, stream VI can be fed to the MTBE plant,preferably into the synthesis section. This applies even when the MTBEplant produces only just the amount of MTBE required for thedissociation and the synthesis section thus provides no further outletsfor high-boiling components. A second possibility is to separate off themajor part of the methanol from the stream VI by distillation andrecirculate the remainder to process step a) (stream VIII in FIG. 2). Afurther possibility is to separate off methanol, the by-products and2-methoxybutane from the stream VI by distillation and recirculate theMTBE which remains to process step b). The latter two possibilities areadvantageous, in particular, for stand-alone plants in which MTBEdelivered from outside is used.

The overhead product V which is obtained in process step c) andpreferably comprises more than 95% by mass of isobutene can be useddirectly as commercial product or be purified further.

Since isobutene forms a minimum boiling point azeotrope with methanol,the overhead product V obtained in process step c) contains not only themain product isobutene but also, in particular, methanol. Furthercomponents which can be present in the overhead product V are, forexample, dimethyl ether, which can be formed, for example, bycondensation of methanol, and linear butenes (1-butene, cis-2-butene,trans-2-butene), which can be formed, for example, by decomposition of2-methoxybutane, and water.

Part of the dimethyl ether can optionally be separated off from theoverhead product V in process step c) by operating the condenser on thedistillation column or reactive distillation column as a partialcondenser. The C₄ fraction present in the overhead product can becondensed in this and part of the dimethyl ether present in the overheadproduct V can be taken off in gaseous form.

The content of linear butenes in the overhead product V obtained inprocess step c) is, based on the C₄-olefin fraction, preferably lessthan 1000 ppm by mass. The content of 1-butene in the overhead fractionV based on the C₄-olefin fraction is preferably less than 500 ppm bymass. The content of 2-butenes (sum of the two 2-butenes) is, based onthe C₄-olefin fraction, preferably likewise less than 500 ppm by mass.

Process Step d) Isobutene Work-Up

Commercial isobutene grades are usually virtually free of methanol. Themethanol can be separated off from the stream V obtained in process stepc) by methods known per se, for example by extraction. The extraction ofmethanol from stream V can be carried out using, for example, water oran aqueous solution as extractant, e.g. in an extraction column. Theextraction is preferably carried out using water or an aqueous solutionin an extraction column which preferably has from 4 to 16 theoreticalplates. The extractant is preferably passed through the extractioncolumn in countercurrent relative to the stream to be extracted. Theextraction is preferably carried out at a temperature of from 15 to 50°C., more preferably from 25 to 40° C. For example, when an extractioncolumn having more than 6 theoretical plates is used and is operated ata pressure of 0.9 MPa_((abs)) and a temperature of 40° C., awater-saturated isobutene having an isobutene content above 99% by masscan be obtained.

The methanol-containing water extract obtained in the extraction can beseparated into water and methanol by distillation. The water can berecirculated as extractant to the extraction stage. The methanol can beutilized for customary industrial syntheses, for example esterificationsor etherifications.

The moist isobutene stream from the extraction column can be worked upin a further distillation column by removal of water and dimethyl etherto give dry isobutene. The dry isobutene is obtained as bottom producthere. In the condensation system at the top of the column, water istaken off in liquid form and dimethyl ether can be taken off in gaseousform after phase separation. A distillation column which is preferablyused for drying preferably has from 30 to 80 theoretical plates,preferably from 40 to 65 theoretical plates. The reflux ratio is,depending on the number of theoretical plates realized and the requiredpurity of the isobutene, preferably less than 60, more preferably lessthan 40. The operating pressure of the column K2 can preferably be setto from 0.1 to 2.0 MPa(abs).

A work-up of isobutene by extraction and distillation is described indetail in, for example, DE 102 38 370.

The isobutene obtained in this way can, for example, have thecomposition shown in Table 1:

TABLE 1 Typical composition of commercial isobutene Proportions by mass[kg/kg] C₃-Hydrocarbons <0.000100 Butanes <0.001000 Isobutene >0.9990001-Butene <0.000500 2-Butenes <0.000500 Methanol <0.000030C₅-Hydrocarbons <0.000500 Water <0.000050

However, depending on purity requirements, lower concentrations of thesecondary components are also conceivable if necessary.

The isobutene prepared by the process of the invention can, for example,be used for the preparation of methallyl chloride, methallylsulphonates,methacrylic acid or methyl methacrylate. In particular, it can beadvantageous for both the methanol and the isobutene to be separated offfrom the dissociation product and for both the methanol and theisobutene to be used for the preparation of methyl methacrylate. Such aprocess for preparing methyl methacrylate is described, for example, inEP 1 254 887, which is expressly incorporated by reference.

Starting Material

MTBE of differing quality can be used as MTBE-containing stream I inprocess step a). In particular, it is possible to use industrial MTBE ofvarious qualities or mixtures of industrial MTBE and methanol as streamI. Industrial MTBE (fuel grade) is the preferred starting material.Table 2 shows, for example, the typical composition of an industrialMTBE from OXENO Olefinchemie GmbH.

TABLE 2 Typical composition of industrial MTBE (fuel grade) from OxenoProportions by mass [kg/kg] 1-Butene 0.000080 2-Butenes 0.000920Pentanes 0.001500 MTBE 0.978000 2-Methoxybutane 0.003000 Methanol0.008500 tert-Butanol 0.003000 Water 0.000050 Diisobutene 0.003300

Industrial MTBE can be prepared by known methods by reactingC₄-hydrocarbon mixtures from which the multiply unsaturated hydrocarbonshave largely been removed, for example raffinate I or selectivelyhydrogenated cracking C₄, with methanol. A process for preparing MTBE isdescribed, for example, in DE 101 02 062.

In the process of the invention, a mixture of industrial MTBE and asubstream of the dissociation product IV which has been separated off bydistillation can also be used as stream I. The substream here can inparticular be the stream obtained in process step c) as bottom productVI or a stream which can be obtained by working up the bottom stream VIobtained in process step c).

It can be particularly advantageous in the process of the invention foran MTBE-containing stream which is entirely or partly obtained byremoving low boilers from an MTBE-containing stream Ia in an optionalprocess step e) to be used as stream I. The substream separated off fromthe dissociation product IV by distillation can be added to this streamI, too.

The removal of low boilers can be particularly advantageous when theMTBE-containing stream Ia comprises, for example, C₄- orC₅-hydrocarbons. The separation of the low boilers from stream Ia in theoptional process step e) can preferably be carried out in a distillationcolumn. The distillation column is preferably operated so that the lowboilers can be separated off as overhead product.

The process step e) is preferably carried out in a distillation columnwhich has from 30 to 75 theoretical plates, preferably from 40 to 65theoretical plates and particularly preferably from 40 to 55 theoreticalplates. The column is, depending on the number of theoretical platesrealized, the composition of the MTBE used and the required purity ofC₄- and C₅-hydrocarbons, preferably operated at a reflux ratio of from150 to 350, in particular from 200 to 300. The column in the optionalprocess step e) is preferably operated at a pressure of from 0.2 to 0.6MPa_((abs)), more preferably from 0.3 to 0.4 MPa_((abs)). The column canbe heated using, for example, steam. The condensation can, depending onthe operating pressure selected, be carried out by means of coolingbrine, cooling water or air. The vapour from the top of the column canbe completely or only partially condensed, so that the overhead productVII can be taken off in either liquid or vapour form. The overheadproduct VII can be utilized thermally or be used as starting materialfor a synthesis gas plant.

If columns such as the columns denoted by K1, K2 and K3 in FIG. 1 orFIG. 2 are used in the process of the invention, these can be providedwith internals such as trays, rotating internals, irregular beds and/orordered packing.

In the case of column trays, the following types, for example, can beused:

-   trays having holes or slits in the tray plate;-   trays having necks or chimneys which are covered by bubble caps,    covers or hoods;-   trays having holes in the tray plate which are covered by movable    valves;-   trays having special constructions.

In columns having rotating internals, the runback can, for example, besprayed by means of rotating funnels or spread as a film on a heatedtube wall by means of a rotor.

As mentioned above, columns having irregular beds of various packingelements can be used in the process of the invention. The packingelements can comprise virtually all materials, in particular, forexample, steel, stainless steel, copper, carbon, stoneware, porcelain,glass or plastics, and have a wide variety of shapes, in particularspheres, rings having smooth or profiled surfaces, rings having internalstruts or openings through the wall, wire meshes, saddle bodies andspirals.

Packings having a regular/ordered geometry can, for example, comprisemetal sheets or meshes. Examples of such packings are Sulzer meshpackings BX made of metal or plastic, Sulzer lamella packings Mellapakmade of sheet metal, Sulzer high-performance packings such asMellapakPlus, structured packings from Sulzer (Optiflow), Montz (BSH)and Kühni (Rombopak).

The present invention is illustrated below with reference to FIGS. 1 and2, without the invention being restricted to the embodiments depicted byway of example there.

A block diagram of an embodiment of a plant in which the process of theinvention can be carried out is shown in FIG. 1. The MTBE-containingfeed stream I separated in the column K1 into a bottom product III whichcontains the secondary components having boiling points higher than thatof MTBE, for example diisobutene and 2-methoxybutane, and anMTBE-containing overhead product II. The MTBE present in the overheadproduct II is mostly dissociated into isobutene and methanol in thereactor R. The dissociation product IV obtained in reactor R isfractionated in the column K2 to give an isobutene-containing mixture Vcontaining methanol and possibly dimethyl ether and a bottom product VIcontaining unreacted MTBE, methanol and secondary components.

A recycle stream VIII from the dissociation, which comprises principallyunreacted MTBE, can optionally be added to the feed stream I. Such arecycle stream VIII can, for example, be the bottom product VI from thecolumn K2, in which case the methanol present therein should be partlyor completely separated off by means of one or more distillation stepsbefore recirculation. The column K2 can optionally be configured as areactive distillation column.

A block diagram of a further embodiment of a plant in which the processof the invention can be carried out is shown in FIG. 2. The embodimentof FIG. 2 differs from that of FIG. 1 in that the feed MTBE Ia isseparated in a preliminary column K3 into an overhead product VIIcomprising low boilers such as C₄- and C₅-hydrocarbons and anMTBE-containing stream I which has been freed of low boilers and is thenfed into the column K1.

Having generally described this invention, a further understanding canbe obtained by reference to certain specific examples which are providedherein for purposes of illustration only, and are not intended to belimiting unless otherwise specified.

EXAMPLES Example a Production of a Shaped Aluminosilicate Body

500 g of aluminosilicate powder (manufacturer: Grace Davison, grade:Davicat O 701, formal Al₂O₃ content: 13% by mass, formal SiO₂ content:76% by mass, formal Na₂O content: 0.1% by mass, loss on ignition at 85°C.: about 11%), 363 g of Disperal gel (formal Al₂O₃ content: 15.6%),which was obtained by stirring 197 g of Disperal, a boehmite having aformal Al₂O₃ content of 77% by mass from Sasol Deutschland GmbH, into803 g of a 1.28% strength by mass aqueous nitric acid, subsequentlystirring vigorously so that the gel being formed was continually shearedand was thus kept in a fluid state in a covered container for 3 hours at60° C., cooling the gel to room temperature and replacing any waterwhich had evaporated, and 370 g of deionized water were firstly mixedthoroughly with one another in a high-intensity mixer from Eirich.Pelletization was subsequently carried out in the high-intensity mixerfrom Eirich, giving uniform roundish pellets having a diameter of fromabout 1 to 3 mm over a period of 30-40 minutes. The moist pellets werefirstly dried at 120° C. in a stream of air and subsequently heated at 2K/min to 550° C. and calcined at this temperature for 10 hours in astream of air. The aluminosilicate pellets produced in this way formallycontained 76% by mass of Al₂O₃ and 24% by mass of SiO₂. In addition, thecatalyst produced contained 0.12% by mass of sodium compounds(calculated as sodium oxide). The composition of the aluminosilicatepellets was calculated from the amount and composition of the startingsubstances. The aluminosilicate pellets had a pore volume determined bythe above-described cyclohexane method of 1.15 ml/g.

Example b Production of a Shaped Catalyst (According to the Invention)

An impregnation solution having a magnesium content of 4.7% by mass wasproduced from deionized water and magnesium nitrate hexahydrate. The pHof this solution was 5.1. A sieve fraction of the aluminosilicatesupport produced in Example 1 (diameter: 1.0 mm-2.8 mm) was impregnatedwith the acidic magnesium nitrate solution by vacuum impregnation. Forthis purpose, the pellets were placed in a glass tube and the latter wasevacuated for about 30 minutes (water pump vacuum of about 25 hPa). Theimpregnation solution was subsequently sucked in from the bottom toabove the upper surface of the bed of solid. After a contact time ofabout 15 minutes, the solution which had not been taken up by thesupport was drained. The moist pellets were firstly dried to constantweight at 140° C. in a stream of air and subsequently heated at 3 K/minto 450° C. and calcined at this temperature for 12 hours. The catalystproduced formally comprised 68% by mass of silicon dioxide, 21% by massof aluminum oxide and 11% by mass of magnesium oxide. In addition, thecatalyst produced contained 0.11% by mass of sodium compounds(calculated as sodium oxide). The composition of the catalyst wascalculated from the amount and composition of the starting substancesand of the impregnation solution which had been drained off. The amountsof sodium were a constituent of the aluminosilicate used in Example 1.The pore volume determined by the above-described cyclohexane method was1.1 ml/g.

The following example calculations were carried out using thesteady-state simulation program ASPEN Plus (version 12.1 fromAspenTech). To produce transparent, reproducible data, only generallyavailable materials data were used. In addition, the use of a reactivedistillation was dispensed with in all variants. These simplificationsmake it readily possible for a person skilled in the art to reproducethe calculations. Although the methods used do not have sufficientaccuracy for the design of industrial plants, the qualitativedifferences between the arrangements were correctly determined. In allvariants demonstrated, the MTBE conversion can be increased by use of areactive distillation.

The Property Method “UNIFAC-DMD” (see J. Gmehling, J. Li, and M.Schiller, Ind. Eng. Chem. Res. 32, (1993), pp. 178-193) was used in theexamples. Modelling was in each case based on a reactor volume of thereactor R of 100 l, assuming a charge of a catalyst which formallycomprises magnesium oxide, aluminum oxide and silicon oxide and whoseproduction was described in Examples a and b.

For modeling of the reactor, a kinetic reactor model based oncomprehensive experimental measured data using this catalyst was used inthe calculations. The reaction temperatures assumed in the modeling ofthe reactor were therefore also reported in the examples. Since thecomposition of the inflowing and outflowing streams of the reactionstage were also reported in each case, a person skilled in the art can,by reproducing the reactors with prescribed conversions, repeat thecalculation of the example without knowing the precise equations for thekinetics. The reactor pressure was 0.8 MPa_((abs)) in all examples.

Example 1

Example 1 corresponds to the variant shown in FIG. 2 without theoptional recirculation of dissociation product to the column K1. An MTBEstream (Ia) of 100 kg/h having the composition shown in Table 3 wasassumed as feed to the MTBE dissociation plant, as shown in FIG. 2. Thecontent of 2-methoxybutane was higher than that of commercial fuel MTBE(see Table 1). This was typical of MTBE from MTBE synthesis plants whichwere coupled with an MTBE dissociation plant. The only outlet for2-methoxybutane in such a coupled plant was the column K1 in which partof the 2-methoxybutane was separated off via the bottom. Therecirculation of streams containing 2-methoxybutane from the MTBEdissociation to the MTBE synthesis therefore results in increases in the2-methoxybutane concentration in the MTBE.

TABLE 3 Composition of the assumed MTBE inflow stream into the MTBEdissociation plant for Example 1 Feed MTBE (Ia) Mass flow [kg/h] 100.00Proportions by mass [kg/kg] Dimethyl ether Isobutene 1-Butene 0.0000802-Butenes 0.000920 Pentanes 0.001500 MTBE 0.978150 2-Methoxybutane0.004500 Methanol 0.008500 tert-Butanol 0.003000 Water 0.000050Diisobutene 0.003300

In the column K3, the C₄- and C₅-hydrocarbons were firstly separatedfrom the MTBE stream (Ia) down to a residual content of 50 ppm by mass.The column had 50 theoretical plates and was operated at a reflux ratioof 210 and a pressure of 0.5 MP_((abs)). The feed of the crude MTBE Iawas introduced above plate 20, counted from the top. The temperature atthe top was 63.0° C., and the temperature at the bottom was 111.9° C.The distillate from this column (VII) had a residual MTBE content of 10%by mass. The MTBE content can be reduced further by increasing thereflux ratio and/or number of theoretical plates. Table 4 shows thecomposition of the distillate stream (VII) and the bottom stream fromthe column K3.

TABLE 4 Composition of the distillate stream (VIII) and of the bottomstream to the column K3 for Example 1 Distillate from K3 Bottom product(VII) from K3 Mass flow [kg/h] 0.31 99.69 Proportions by mass [kg/kg]Dimethyl ether Isobutene 1-Butene 0.025943 2-Butenes 0.298339 Pentanes0.470265 0.000050 MTBE 0.100052 0.980866 2-Methoxybutane 0.0000230.004514 Methanol 0.089236 0.008250 tert-Butanol 0.003009 Water 0.016143Diisobutene 0.003310

The MTBE stream I (bottom product from column K3) which had largely beenfreed of low boilers was fed to the column K_(i) in which principallydiisobutene and 2-methoxybutane were separated off via the bottom (III).The column had 90 theoretical plates and was operated at a reflux ratioof 4.6 and a pressure of 0.9 MPa_((abs)). The feed stream I wasintroduced above plate 30, counted from the top. The temperature at thetop was 141.3° C., and the temperature at the bottom was 145.8° C. Agaseous fraction which contains 98.5% by mass of MTBE was obtained asoverhead product (II). The 2-methoxybutane content of the distillate wasset to 2000 ppm by mass (see Table 5). The MTBE content of the bottomproduct can be reduced by increasing the reflux ratio and/or theseparation power.

TABLE 5 Composition of the distillate stream (II) and bottom stream(III) from column K1 for Example 1 Bottom product Distillate from K1from K1 (II) (III) Mass flow [kg/h] 96.19 3.50 Proportions by mass[kg/kg] Dimethyl ether Isobutene 1-Butene/2-butenes Pentanes 0.000052MTBE 0.986281 0.832050 2-Methoxybutane 0.002000 0.073603 Methanol0.008550 tert-Butanol 0.003117 0.000062 Water Diisobutene 0.094286

The distillate stream (II) from the column K1 is, after further heatingto the reaction temperature, fed to the reaction section (R). Thereactor was operated at 300° C. and 0.8 MPa_((abs)). Under thesereaction conditions, an MTBE conversion of about 96% was obtained andthe conversion of 2-methoxybutane was about 22%. However, since thecontent of 2-methoxybutane in the reactor feed was limited to 2000 ppmby mass, the required specification for linear butenes in the isobuteneproduct was not put at risk in spite of the dissociation of2-methoxybutane to 2-butene. The composition of the reactor output (IV)is shown in Table 6.

TABLE 6 Composition of the reactor output (IV) and of the distillatestream (V) and the bottom stream (VI) from column K2 for Example 1Reactor Distillate Bottom product output from K2 from K2 (IV) (V) (VI)Mass flow [kg/h] 96.19 60.57 35.63 Proportions by mass [kg/kg] Dimethylether 0.003430 0.005264 0.000311 Isobutene 0.604965 0.960218 0.0010001-Butene/2-butenes 0.000284 0.000449 0.000003 Pentanes 0.000052 0.0000760.000011 MTBE 0.037479 0.101194 2-Methoxybutane 0.001554 0.004196Methanol 0.348828 0.033809 0.884392 tert-Butanol 0.000773 0.002087 Water0.001911 0.000183 0.004849 Diisobutene 0.000725 0.001957

The reactor output (IV) was partially condensed and fed as a two-phasemixture to the column K2. The column had 40 theoretical plates and wasoperated at a reflux ratio of 0.3 and a pressure of 0.6 MPa_((abs)). Thefeed stream was introduced above plate 30, counted from the top. Thetemperature at the top was 48.4° C., and the temperature at the bottomwas 114.8° C. The bottom product comprises predominantly unreacted MTBE(about 10% by mass) and methanol (about 88% by mass), see Table 6. Thisbottom product can be recirculated to an MTBE synthesis plant. Sincethere was usually no outlet for 2-methoxybutane in the synthesis plant,the 2-methoxybutane content of the MTBE therefore increases, as takeninto account in this example, compared to MTBE produced in an MTBEsynthesis without coupling with an MTBE dissociation, see Tables 2 and3.

The overhead product (V) was isobutene having a purity of greater than96% by mass of isobutene. The limits for linear butenes required in atypical isobutene specification (<1000 ppm by mass) and those forC₅-hydrocarbons (<1000 ppm by mass) were reliably adhered to, cf. Tables1 and 6). If required, the methanol can be removed by extraction withwater, and the residual water and the dimethyl ether can be removed bymeans of a subsequent distillation and the isobutene can be concentratedto a purity of greater than 99.9% by mass. An isobutene which had beenpurified further in this way meets the specification given in Table 2.

Example 2

Example 2 corresponds to the variant shown in FIG. 2 with recirculationof dissociation product to the column K1. According to FIG. 2, an MTBEstream (Ia) of 100 kg/h having the composition shown in Table 7 (typicalfuel MTBE, compare with Table 2) was assumed as feed to the MTBEdissociation plant.

TABLE 7 Composition of the assumed MTBE inflow stream into the MTBEdissociation plant for Example 2 Feed MTBE (Ia) Mass flow [kg/h] 100.00Proportions by mass [kg/kg] Dimethyl ether Isobutene 1-Butene 0.0000802-Butenes 0.000920 Pentanes 0.001500 MTBE 0.979650 2-Methoxybutane0.003000 Methanol 0.008500 tert-Butanol 0.003000 Water 0.000050Diisobutene 0.003300

In the column K3, the low-boilers (C₄- and C₅-hydrocarbons) were onceagain firstly separated off from the MTBE stream (Ia) down to a residualcontent of 50 ppm by mass. Number of theoretical plates, point ofintroduction of the feed, reflux ratio and operating pressure of thecolumn K3 were unchanged from Example 1. The temperature at the top was63.0° C., and the temperature at the bottom was 111.8° C. The distillatefrom this column (VII) had a residual MTBE content of 10% by mass. TheMTBE content can be reduced further by increasing the reflux ratioand/or number of theoretical plates. Table 8 shows the composition ofthe distillate stream (VII) and the bottom stream from column K3.

TABLE 8 Composition of the distillate stream (VIII) and the bottomstream from column K3 for Example 2 Distillate from K3 Bottom product(VIII) from K3 Mass flow [kg/h] 0.31 99.69 Proportions by mass [kg/kg]Dimethyl ether Isobutene 1-Butene 0.025947 2-Butenes 0.298386 Pentanes0.470332 0.000050 MTBE 0.100001 0.982371 2-Methoxybutane 0.0000150.003009 Methanol 0.089175 0.008250 tert-Butanol 0.003009 Water 0.016145Diisobutene 0.003310

The MTBE which had largely been freed of low boilers was admixed with arecycle stream (VIII) from the MTBE dissociation to form the stream (I).This was the bottom stream (VI) from column K2, it being assumed that amajor part of the methanol and the secondary components had been removedby one or more distillation steps prior to recirculation. The assumedcomposition of the recycle stream (VIII) and of the feed stream (I) tothe column K1 formed by this mixing is shown in Table 9.

TABLE 9 Composition of the recycle stream from the dissociation (VIII)and the feed stream (I) to the column for Example 2 Recycle stream Feedto K1 (VIII) (I) Mass flow [kg/h] 22.99 122.69 Proportions by mass[kg/kg] Dimethyl ether Isobutene 1-Butene 2-Butenes Pentanes 0.000041MTBE 0.771297 0.942811 2-Methoxybutane 0.011253 0.004554 Methanol0.217450 0.047459 tert-Butanol 0.002445 Water Diisobutene 0.002690

The stream (I) formed from the bottom product from column K3 and therecycle stream (VIII) from the MTBE dissociation by mixing was fed tocolumn K1 in which principally diisobutene and 2-methoxybutane wereagain removed via the bottom (III). Number of theoretical plates, pointof introduction of the feed stream and operating pressure of column K1were unchanged from Example 1. The column was operated in Example 2 at areflux ratio of 3.2. The temperature at the top was 135.8° C., and thetemperature at the bottom was 147.2° C. The overhead product obtained(II) was a gaseous fraction which contains about 94.5% by mass of MTBE.The 2-methoxybutane content of the distillate was set to 2500 ppm bymass (cf. Table 10). The MTBE content of the bottom product can bereduced by increasing the reflux ratio and/or the separation power.

TABLE 10 Composition of the distillate stream (II) and bottom stream(III) from column K1 for Example 2 Bottom product Distillate from K1from K1 (II) (III) Mass flow [kg/h] 120.79 1.90 Proportions by mass[kg/kg] Dimethyl ether Isobutene 1-Butene 2-Butenes Pentanes 0.000041MTBE 0.946770 0.691130 2-Methoxybutane 0.002500 0.135146 Methanol0.048205 tert-Butanol 0.002483 0.000040 Water Diisobutene 0.173684

The distillate stream from column K1 is, after further heating to thereaction temperature, fed to the reaction section (R). The reactor wasoperated at 265° C. and 0.8 MPa_((abs)). Under these conditions, an MTBEconversion of about 85% was obtained and the conversion of2-methoxybutane was about 15%. However, since the 2-methoxybutane in thereactor feed was limited to 2500 ppm by mass, the required specificationfor linear butenes in the isobutene product was not put at risk in spiteof the dissociation of 2-methoxybutane to linear butenes. Thecomposition of the reactor output (IV) is shown in Table 11.

TABLE 11 Composition of the reactor output (IV) and of the distillatestream (V) and the bottom stream (VI) of column K2 for Example 2 ReactorDistillate Bottom product output from K2 from K2 (IV) (V) (VI) Mass flow[kg/h] 120.79 64.09 56.70 Proportions by mass [kg/kg] Dimethyl ether0.002742 0.005144 0.000026 Isobutene 0.510121 0.960544 0.0010001-Butene/2-butenes 0.000241 0.000451 0.000003 Pentanes 0.000041 0.0000630.000017 MTBE 0.146749 0.312622 2-Methoxybutane 0.002122 0.004521Methanol 0.335334 0.033554 0.676443 tert-Butanol 0.000546 0.001164 Water0.001543 0.000243 0.003012 Diisobutene 0.000560 0.001193

The reactor output (IV) was partially condensed and fed as a two-phasemixture to the column K2. Number of theoretical plates, point ofintroduction of the feed stream, reflux ratio and operating pressure ofthe column K2 were unchanged from Example 1. The temperature at the topwas 48.3° C., and the temperature at the bottom was 111.4° C. The bottomproduct once again comprises predominantly unreacted MTBE (about 31% bymass) and methanol (about 67% by mass), see Table 11. Owing to thesignificantly higher residual MTBE content compared to Example 1, it wasworthwhile in this example to separate off the major part of themethanol and further secondary components by means of one or moredistillation steps and recirculate the stream which had been purified inthis way to a point upstream of the column K1. The total yield ofisobutene, based on the MTBE used, can be increased significantly inthis way.

The overhead product (V) was isobutene having a purity of greater than96% by mass of isobutene. The limits for linear butenes required in atypical isobutene specification (<1000 ppm by mass) and those forC₅-hydrocarbons (<1000 ppm by mass) were reliably adhered to, cf. Tables1 and 11. If required, the methanol can be removed by extraction ofwater, the residual water and the dimethyl ether can be separated off bymeans of a subsequent distillation and the isobutene can be concentratedto a purity of greater than 99.9% by mass. An isobutene which had beenpurified further in this way meets the specification given in Table 1.

German patent application 10 2006 040430.0 filed Aug. 29, 2006, isincorporated herein by reference.

Numerous modifications and variations on the present invention arepossible in light of the above teachings. It is therefore to beunderstood that within the scope of the appended claims, the inventionmay be practiced otherwise than as specifically described herein.

1. A process for the dissociation of methyl tert-butyl ether (MTBE),comprising: a) separating an MTBE-containing stream I by distillationinto an MTBE-containing overhead stream II and a bottom stream III whichcomprises compounds having boiling points higher than that of MTBE; b)dissociation of the MTBE present in the overhead stream II over acatalyst to give a dissociation product IV; c) separating thedissociation product IV by distillation into an isobutene-containingoverhead stream V and a bottom stream VI comprising unreacted MTBE; andd) separating off methanol from the isobutene-containing overhead streamV obtained in process step c) by extraction wherein dimethyl ether andwater are separated off by distillation from the extracted isobutene;wherein the stream I has a proportion of 2-methoxybutane (MSBE) ofgreater than 1000 ppm by mass, based on MTBE, and wherein the separationby distillation in step a) and/or the dissociation in step b) is carriedout so that the dissociation product IV has a concentration of less than1000 ppm by mass of linear butenes, based on a C₄-olefin fraction. 2.The process according to claim 1, wherein the distillation in step a) isa fractional distillation which is carried out so that the stream II hasa concentration of 2-methoxybutane based on MTBE of less than 1000 ppmby mass.
 3. The process according to claim 1, wherein the dissociationin step b) is carried out under conditions under which a conversion ofMTBE is greater than a conversion of 2-methoxybutane.
 4. The processaccording to claim 1, wherein the dissociation in step b) is carried outover a catalyst which has an activity in respect of the dissociation ofMTBE which is at least 1% greater than the activity in respect of thedissociation of 2-methoxybutane.
 5. The process according to claim 1,wherein the dissociation in step b) is carried out at a conversion ofMTBE of at least 70%.
 6. The process according to claim 1, whereinindustrial MTBE is used as stream I.
 7. The process according to claim1, wherein a mixture of industrial MTBE and a substream separated offfrom the dissociation product IV by distillation is used as stream I. 8.The process according to claim 1, wherein an MTBE-containing streamwhich is entirely or partly obtained by removing low boilers from anMTBE-containing stream Ia is used as stream I.
 9. The process accordingto claim 1, wherein the separation by distillation in step a) is carriedout so that the dissociation product IV has a concentration of less than1000 ppm by mass of linear butenes, based on a C₄-olefin fraction. 10.The process according to claim 1, wherein the dissociation in step b) iscarried out so that the dissociation product IV has a concentration ofless than 1000 ppm by mass of linear butenes, based on a C₄-olefinfraction.
 11. The process according to claim 1, wherein the separationby distillation in step a) and the dissociation in step b) are carriedout so that the dissociation product IV has a concentration of less than1000 ppm by mass of linear butenes, based on a C₄-olefin fraction. 12.The process according to claim 1, wherein the dissociation product IVcomprises isobutene.
 13. The process according to claim 1, wherein thestream I has a proportion of 2-methoxybutane of greater than 3,000 ppmby mass.
 14. The process according to claim 1, wherein the overheadstream II has a proportion of 2-methoxybutane that is less than theproportion of the 2-methoxybutane in the bottom stream III.
 15. Theprocess according to claim 1, wherein the stream I has a proportion of2-methoxybutane of greater than 4,500 ppm by mass.
 16. The processaccording to claim 1, wherein the MTBE-containing stream I is distilledonly once to form the MTBE-containing overhead stream II and the bottomstream III.
 17. The process according to claim 1, wherein the overheadstream II is flowed directly to the dissociating and the overhead streamhas the same composition as the stream subjected to the dissociation.18. The process according to claim 1, wherein the overhead stream IIcomprises 2-methoxybutane and, during the dissociation of the MTBE, the2-methoxybutane is at least partially dissociated to form one or morebutenes.